Fluidized bed gas distributor, reactor using fluidized bed gas distributor, and method for producing para-xylene and co-producing light olefins

ABSTRACT

Disclosed are a fluidized bed gas distributor and a fluidized bed reactor, the fluidized bed reactor comprising a first distributor (1) and a second distributor (2), wherein the first distributor (1) is located at the bottom of a fluidized bed, and second distributor (2) is located downstream of a gas from the first distributor (1). Also disclosed is a method for producing a para-xylene and co-producing light olefins, the method comprising the following steps: material stream A enters a reaction zone (3) of a fluidized bed reactor from a first gas distributor (1); material stream B enters the reaction zone (3) of the fluidized bed reactor from a second gas distributor (2); and the reactants are brought into contact with a catalyst in the reaction zone (3) to generate a gas phase stream comprising para-xylene and light olefins.

TECHNICAL FIELD

The present invention relates to a distributor in a fluidized bed, areactor and a production method for producing para-xylene (PX) andco-producing light olefins, and is particularly suitable for a fluidizedbed reactor and a production method for preparing para-xylene andco-producing light olefins by the alkylation of toluene and methanol,which belongs to the field of chemistry and chemical industry.

BACKGROUND

Para-xylene (PX) is one of the basic organic raw materials in thepetrochemical industry, which has a wide range of applications inchemical fiber, synthetic resins, pesticides, pharmaceuticals andpolymer materials. At present, the production of para-xylene mainly usestoluene, C₉ aromatics and mixed xylene as raw materials, and para-xyleneis obtained by disproportionation, isomerization, adsorption separationor cryogenic separation. Since the para-xylene content in the product iscontrolled by thermodynamics, para-xylene only accounts for 22˜24% ofthe C₈ mixed aromatics, and the material circulation processing amountis large during the process, and the equipment is large and theoperation cost is high. In particular, the three isomers of xylene havesmall differences in boiling points, and it is difficult to obtainhigh-purity para-xylene by conventional distillation techniques, and anexpensive adsorption separation process must be employed. In recentyears, many patents, domestic and abroad, have disclosed a new route forthe production of para-xylene. The toluene-methanol alkylationtechnology is a new way to produce para-xylene with high selectivity,which has been highly valued and paid great attention by the industry.

Light olefins, namely ethylene and propylene, are two basicpetrochemical feedstocks that are increasingly in demand. Ethylene andpropylene are mainly produced from naphtha, depending on the petroleumroute. In recent years, the non-petroleum route to produce ethylene andpropylene has received more and more attention, especially the processroute of the methanol conversion to light olefins (MTO), which is animportant way to achieve petroleum substitution strategy, reduce andalleviate our demand and dependence for petroleum.

The above mentioned new routes for the preparation of para-xylene andlight olefins are all acid-catalyzed reaction. Methanol-to-olefinsreaction is inevitable in the process of preparing para-xylene by thealkylation of toluene and methanol based on the ZSM-5 molecular sievecatalyst. In the course of this reaction, the following reactions occurmainly:

C₆H₅—CH₃+CH₃OH→C₆H₄—(CH₃)₂+H₂O  (1)

nCH₃OH→(CH₂)n+nH₂O n=2, 3  (2)

A conventional toluene alkylation process involves mixing toluene andmethanol upstream of the reactor and then feeding the mixture togetherinto the reactor. The reactor type includes a fixed bed and a fluidizedbed. In order to increase the conversion rate of toluene, the phasedinjection of reactants has been employed in various fixed bed andfluidized bed processes.

The competition between the MTO reaction and the alkylation reaction isa major factor affecting the conversion rate of toluene and the yield ofthe para-xylene. The process of simultaneously realizing two reactionsin the same reactor is simple, but the conversion rate of toluene islow; the process of respectively realizing two reactions in differentreactors is complicated, but the conversion rate of toluene and theyield of para-xylene are higher. Therefore, the process of thealkylation of toluene and methanol to prepare para-xylene and co-producelight olefins requires a major breakthrough in the process configurationand the reactor design, thereby coordinating and optimizing thecompetition between the alkylation reaction and the MTO reaction, andimproving the conversion rate of toluene and the yield of para-xylene.

SUMMARY OF THE INVENTION

According to an aspect of the present application, this is provided afluidized bed gas distributor to achieve mass transfer control bydistributing different raw materials stream in different regions in aco-feed system with a large difference in raw material reaction rates,so as to coordinate and optimize a co-feed system and improve thereaction yield. As a typical reaction system, the reaction of preparingpara-xylene by the alkylation of toluene and methanol, in which thereaction rates of the alkylation reaction and the MTO reaction aregreatly different, and the MTO reaction inhibits the alkylationreaction. Therefore, the conversion rate of toluene is low. Thefluidized bed gas distributor provided by the present applicationcoordinates and optimizes the competition between the alkylationreaction and the MTO reaction through mass transfer control, therebyimproving the conversion rate of toluene and the yield of para-xylene.

Methanol is both a raw material for the alkylation reaction of tolueneand methanol, and a raw material for the MTO reaction, but the reactionrate of the MTO is much higher than that of the alkylation reaction oftoluene and methanol. As our experimental studies show, when co-feedingof toluene and methanol, the MTO reaction quickly consumes most of themethanol, inhibits the alkylation reaction of toluene and methanol, andthe conversion rate of toluene is low. It can be seen from the aboveanalysis that the technical field needs to coordinate and optimize thecompetition between the alkylation reaction and the MTO reaction fromthe two aspects of the catalyst design and the reactor design.

The fluidized bed gas distributor of the present application comprises afirst distributor and a second distributor, wherein the firstdistributor is located at the bottom of the fluidized bed, and thesecond distributor is located in at least one region of the gas flowdownstream from the first distributor.

Preferably, the second distributor is a microporous gas distributor.

Preferably, the second distributor comprises an intake pipe, amicroporous pipe and an intake ring pipe;

the intake pipe is connected with a gas path of the microporous pipe,the gas is introduced by the intake pipe from the outside of thefluidized bed into the microporous pipe in the fluidized bed;

the intake ring pipe is connected with a gas path of the intake pipe,the intake ring pipe is disposed on a plane perpendicular to the flowdirection of the gas from the first distributor;

the microporous pipe is disposed on the intake ring pipe andperpendicular to a plane of the intake ring pipe.

Preferably, material stream A enters the fluidized bed through the firstdistributor, material stream B enters the fluidized bed through thesecond distributor and contacts with at least a portion of the gas ofthe material stream A.

As a preferred embodiment, the fluidized bed of the present applicationis a fluidized bed reactor for producing para-xylene and co-producinglight olefins from methanol and toluene.

Preferably, the first distributor is a two-dimensional gas distributorand the first distributor distributes the gas on the plane in which thefirst distributor is located at the bottom of the fluidized bed.

Preferably, the second distributor is a three-dimensional gasdistributor and the second distributor distributes the gas in at least aportion of the reaction space in which the second distributor is locatedin the fluidized bed.

In the present application, “at least a portion of the reaction space”refers to at least a portion of the space within the reaction zone.

Preferably, the first distributor is a branched pipe distributor and/ora plate distributor with blast caps.

Preferably, the microporous pipe is a ceramic microporous pipe and/or ametal microporous sintered pipe.

Preferably, the side and end faces of the microporous pipe havemicropores with a pore diameter ranging from 0.5 μm to 50 μm.

Preferably, the side and end faces of the microporous pipe havemicropores with a porosity ranging from 25% to 50%.

Preferably, the gas velocity in the pipe of the microporous pipe is in arange from 0.1 m/s to 10 m/s.

Preferably, the gas velocity in the pipe of the microporous pipe is in arange from 1 m/s to 10 m/s.

Preferably, the microporous pipes are arranged in plurality and arrangedin parallel with each other; the intake ring pipes are arranged inplurality and arranged in a concentric ring or planar spiral in the sameplane.

Preferably, the fluidized bed gas distributor is used in a fluidized bedreactor for producing para-xylene and co-producing light olefins.

According to another aspect of the present application, there isprovided a fluidized bed reactor, which is used for a co-feed systemwith a large difference in raw material reaction rates. It candistribute different raw materials stream in different regions toachieve mass transfer control, so as to coordinate and optimize aco-feed system and improve the reaction yield. As a typical reactionsystem, the reaction of preparing para-xylene by the alkylation oftoluene and methanol, in which the reaction rates of the alkylationreaction and the MTO reaction are greatly different, and the MTOreaction inhibits the alkylation reaction. Therefore, the conversionrate of toluene is low. The fluidized bed reactor provided by thepresent application coordinates and optimizes the competition betweenthe alkylation reaction and the MTO reaction through mass transfercontrol, thereby improving the conversion rate of toluene and the yieldof para-xylene.

The fluidized bed reactor provided by the present application comprisesat least one of the fluidized bed gas distributors of the above aspect.

Preferably, the fluidized bed reactor comprises a reaction zone, asettling zone, a gas-solid separator, a stripping zone and a regeneratedcatalyst delivery pipe;

the first distributor is placed at the bottom of the reaction zone, thesecond distributor is placed above the first distributor, the settlingzone is above the reaction zone, the settling zone is provided with thegas-solid separator, the stripping zone is below the reaction zone, andthe regenerated catalyst delivery pipe is connected with the reactionzone.

As an embodiment, the regenerated catalyst delivery pipe is connectedwith the upper portion of the reaction zone.

As an embodiment, the regenerated catalyst delivery pipe is connectedwith the bottom of the reaction zone.

According to still another aspect of the present application, there isprovided a process for producing para-xylene and co-producing lightolefins from methanol and/or dimethyl ether and toluene. By distributingdifferent raw materials stream in different regions, to achieve masstransfer control, so as to coordinate and optimize a co-feed system andimprove the reaction yield. The reaction of preparing para-xylene by thealkylation of toluene and methanol, in which the reaction rates of thealkylation reaction and the MTO reaction are greatly different, and theMTO reaction inhibits the alkylation reaction. Therefore, the conversionrate of toluene is low. The fluidized bed reactor provided by thepresent application coordinates and optimizes the competition betweenthe alkylation reaction and the MTO reaction through mass transfercontrol, thereby improving the conversion rate of toluene and the yieldof para-xylene.

The method for producing para-xylene and co-producing light olefins frommethanol and toluene, provided by the present application, at least oneof any of the above fluidized bed gas distributors and/or at least oneof any of the above fluidized bed reactors are used; the method forproducing para-xylene and co-producing light olefins comprises at leastthe following steps:

(1) passing the material stream A from the first distributor into thereaction zone of the fluidized bed reactor, the reaction zone containinga catalyst; the material stream A containing toluene, or the materialstream A containing methanol and/or dimethyl ether and toluene;

(2) passing the material stream B containing methanol and/or dimethylether from the second distributor into the reaction zone of thefluidized bed reactor;

(3) in the reaction zone, contacting methanol and/or dimethyl ether andtoluene from the material stream A and/or the material stream B, withthe catalyst to form material stream C comprising para-xylene and lightolefins.

Preferably, the method for producing para-xylene and co-producing lightolefins from methanol and/or dimethyl ether and toluene furthercomprises the following steps:

(4) passing the material stream C into a settling zone and a gas-solidseparator to separate the material stream C to obtain light olefins,para-xylene, chain hydrocarbon by-products, aromatic by-products andunconverted toluene, unconverted methanol and/or dimethyl ether;

(5) returning unconverted methanol and/or dimethyl ether to thefluidized bed reactor via the second distributor; returning the aromaticby-products and unconverted toluene to the fluidized bed reactor via thefirst distributor;

(6) forming a spent catalyst from the catalyst after carbon depositionin the reaction zone, subjecting the spent catalyst to stripping andregenerating in a regenerator to obtain a regenerated catalyst; passingthe regenerated catalyst into the fluidized bed reactor via theregenerated catalyst delivery pipe.

Wherein, the chain hydrocarbon by-product comprises at least one ofmethane, ethane, propane, butane and C₅₊ chain hydrocarbon. The aromaticby-product comprises at least one of benzene, ethylbenzene, o-xylene,m-xylene and C₉₊ arene.

In the present application, low-carbon olefin includes at least one ofethylene, propylene and butene.

In the present application, “methanol and/or dimethyl ether” means thatmethanol in the feedstock may be replaced in whole or in part bydimethyl ether, including three cases: only methanol; or only dimethylether; or both methanol and dimethyl ether.

In the present application, “methanol and/or dimethyl ether and toluene”includes three cases: methanol and toluene; or dimethyl ether andtoluene; or methanol, dimethyl ether and toluene.

Unless otherwise specified, the methanol in the present application maybe replaced by all or part of dimethyl ether and the amount of methanolmay be calculated by converting dimethyl ether into methanol having thesame number of carbon atoms.

Preferably, the mass ratio of methanol in material stream B enteringfrom the second distributor to methanol in material stream A enteringfrom the first distributor is in a range from 1:1 to 10:1. The massratio of methanol here is compared by converting dimethyl ether (ifcontained) into methanol of the same number of carbon atoms.

Preferably, the sum of the mass percentages of methanol and dimethylether in material stream A is in a range from 0% to 30%.

That is, the material stream A entering from the first distributor doesnot contain methanol and/or dimethyl ether, or the sum of the masspercentages of methanol and dimethyl ether in the material stream A doesnot exceed 30%.

Preferably, the sum of the mass percentages of methanol and dimethylether in material stream A is in a range from 2% to 20%.

Preferably, the fluidized bed reactor has a gas phase linear velocityranging from 0.2 m/s to 2 m/s and a reaction temperature ranging from300° C. to 600° C.

Preferably, the regenerator has a gas phase linear velocity ranging from0.2 m/s to 2 m/s and a regeneration temperature ranging from 500° C. to800° C.

The inventors of the present application have found through researchthat the alkylation reaction and the MTO reaction are mass transfercontrol reactions. That is, the mass transfer rate of gas molecules inthe catalyst pores controls the reaction rate, and the mass transferrate of methanol in the catalyst pores is much larger than that oftoluene, the MTO reaction is limited by the mass transfer rate ofmethanol in the catalyst pores, while the alkylation reaction is limitedby the mass transfer rate of toluene in the catalyst pores. Usingtoluene and methanol co-feeding, along the axial direction of thereactor, from upstream to downstream, the concentration of methanoldecreases rapidly and approaches zero, while the concentration oftoluene slowly decreases. In the upstream region of the reactor, thealkylation reaction rate is limited by the mass transfer rate of toluenein the catalyst pores, while in the downstream region of the reactor,with the rapid consumption of methanol and the rapid decrease ofmethanol diffusion, the alkylation reaction rate is limited by the masstransfer rate of methanol in the catalyst pores. In general, theconversion rate of toluene is in a range from 15% to 40% when themixture is fed simultaneously. From the above analysis, maintaining arelatively stable concentration of methanol in the reactor is one of theeffective ways to promote the alkylation reaction.

The present application coordinates and optimizes the competitionbetween the alkylation reaction and the MTO reaction by controlling theconcentrations of methanol and/or dimethyl ether relative to toluenefrom the viewpoint of reactor design and process configuration, andimproves the yield of para-xylene and the selectivity of light olefinsto ensure that neither the situation of the inhibition of the alkylationreaction occurs due to the rapid consumption of most methanol and/ordimethyl ether by the MTO reaction, nor the situation against thealkylation reaction occurs due to far excess content of methanol and/ordimethyl ether, a large number of the MTO reaction occur, and loweradsorption amount of toluene in the catalyst per unit time.

The benefits brought out by the present application include:

(1) this is provided a fluidized bed gas distributor to achieve masstransfer control by distributing different raw materials stream indifferent regions in a co-feed system with a large difference in rawmaterial reaction rates, so as to coordinate and optimize a co-feedsystem and improve the reaction yield.

(2) this is provided a fluidized bed reactor, the above fluidized bedgas distributor is adopted to achieve mass transfer control bydistributing different raw materials stream in different regions andselective recycling in a co-feed system with a large difference in rawmaterial reaction rates, so as to coordinate and optimize a co-feedsystem and improve the reaction yield.

(3) The method for producing para-xylene and co-producing light olefinsfrom methanol and toluene, provided by the present application, hashigher conversion rate of toluene and the selectivity of para-xylene,the conversion rate of toluene is more than 40%, the selectivity ofpara-xylene in the xylene isomer in the product is greater than 95%, themass single-pass yield of para-xylene based on aromatics is greater than38%, the conversion rate of methanol is greater than 90%, and theselectivity of light olefins (ethylene+propylene+butene) in C₁˜C₆ chainhydrocarbon component is greater than 70%, and good technical resultshave been achieved.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic view of a fluidized bed reactor according to anembodiment of the present application.

FIG. 2 is a schematic view of a fluidized bed reactor according to anembodiment of the present application.

FIG. 3 is a side view of microporous gas distributor in the reactionzone according to an embodiment of the present application.

FIG. 4 is a top view of microporous gas distributor in the reaction zoneaccording to an embodiment of the present application.

The reference numerals in the figures are listed as follows:

-   -   1—first gas distributor, 2—second gas distributor, 3—reaction        zone, 4—settling zone, 5—gas-solid separator, 6—stripping zone,        7—regenerated catalyst delivery pipe.    -   2-1—intake pipe, 2-2—intake ring pipe, 2-3—microporous pipe.

DETAILED DESCRIPTION OF THE EMBODIMENT

The present application will be described in detail below with referenceto the embodiments, but the application is not limited to theseembodiments.

Unless otherwise specified, the raw materials and catalysts in theembodiments of the present application are commercially available.

According to one embodiment of the present application, a fluidized bedreactor for producing para-xylene and co-producing light olefins fromtoluene and methanol is shown in FIGS. 1 and 2, and comprises the firstgas distributor 1 located at the bottom of the fluidized bed, the secondgas distributor 2 located at the reaction zone, the reaction zone 3, thesettling zone 4, the gas-solid separator 5, the stripping zone 6 and theregenerated catalyst delivery pipe 7.

The first gas distributor 1 is placed at the bottom of the reaction zone3, the second gas distributor 2 is placed above the first gasdistributor 1, the settling zone 4 is above the reaction zone 3, thegas-solid separator 5 is disposed within the settling zone 4, and theproduct outlet is set on the top. The stripping zone 6 is below thereaction zone 3, and the regenerated catalyst delivery pipe 7 isconnected with the upper or bottom of the reaction zone 3. Theregenerated catalyst enters the reaction zone from the regeneratedcatalyst delivery line 7 and the spent catalyst passes through thestripping zone 6 and enters the regenerator for regeneration.

As an embodiment of the present application, the first gas distributor 1may be a branched pipe distributor.

As an embodiment of the present application, the first gas distributor 1may be one of the plate distributor with blast caps.

As an embodiment of the present application, the second gas distributor2 is a microporous gas distributor. As shown in FIG. 3, the microporousgas distributor includes the intake pipe 2-1, a plurality of intake ringpipes 2-2. The intake ring pipes 2-2 are centered on the axis of thereactor, and a plurality of microporous pipes 2-3 are uniformlydistributed on the intake ring pipes 2-2. A plurality of microporouspipes 2-3 are uniformly distributed. The gas enters the microporouspipes 2-3 through the intake pipe 2-1 and the intake ring pipes 2-2, andone end of the microporous pipe 2-3 is connected with the intake ringpipe 2-2, and the other end is closed. The gas is discharged through themicropores in the microporous pipes 2-3.

The microporous pipe 2-3 can be a ceramic microporous pipe, a powdermetallurgy microporous pipe, and the spacing between microporous pipes2-3 is greater than 50 mm.

As shown in FIG. 3 and FIG. 4, in one embodiment of the presentapplication, there are a total of 12 microporous pipes 2-3, all of whichare arranged on the intake ring pipes 2-2 and perpendicular to the planeof the ring pipe, in a longitudinal parallel arrangement.

The side and end faces of the microporous pipes 2-3 have a uniformmicroporous structure, the pore diameter of the micropores is in a rangefrom 0.5 μm to 50 μm, the porosity is in a range from 25% to 50%, andthe gas velocity in the pipe is in a range from 0.1 m/s to 10 m/s.Preferably, the gas velocity in the pipe is in a range from 1 m/s to 10m/s.

As an embodiment of the present application, the microporous pipe 2-3 isplaced in the reaction zone 3, which can inhibit the growth of bubbles,reduce the back mixing of gas, increase the exchange of substancesbetween the dense phase and the dilute phase, and improve the reactionrate.

As an embodiment of the present application, the catalyst employed is aZSM-5 molecular sieve catalyst.

Due to using toluene and methanol co-feeding, along the axial directionof the reactor, from upstream to downstream, the concentration ofmethanol and/or dimethyl ether decreases rapidly and approaches zero,while the concentration of toluene slowly decreases. In the upstreamregion of the reactor, the alkylation reaction rate is limited by themass transfer rate of toluene in the catalyst pores, while in thedownstream region of the reactor, with the rapid consumption of methanoland the rapid decrease of methanol diffusion, the alkylation reactionrate is limited by the mass transfer rate of methanol in the catalystpores. Maintaining a relatively stable concentration of methanol in thereactor is one of the effective ways to promote the alkylation reaction.

As an embodiment of the present application, the first gas distributor 1belongs to a two-dimensional gas distributor, that is, the raw gas isrelatively uniformly distributed in the plane in which the first gasdistributor 1 is located.

As an embodiment of the present application, the second gas distributor2 (microporous gas distributor) belongs to a three-dimensional gasdistributor, that is, the raw gas is relatively uniformly distributed inthe three-dimensional space in which the second gas distributor 2 islocated.

As an embodiment of the present application, toluene is introduced fromthe first gas distributor 1, and as the reaction proceeds, theconcentration of toluene gradually decreases from upstream to downstreamalong the direction of the reactor axis.

As an embodiment of the present application, a portion of the methanoland/or dimethyl ether is introduced by the first gas distributor 1 andanother portion of the methanol and/or dimethyl ether is introduced bythe second gas distributor 2, which are distributed to the reaction zone3 around the micropore core pipe 2-3 through the micropores denselyarranged on microporous pipe 2-3. Therefore, in the region where thesecond gas distributor 2 is located, the concentration of methanol issubstantially stabilized, and only in the downstream region of thereaction zone 3, the concentration of methanol rapidly decreases. Themethanol diffusion driving force in the region where the second gasdistributor 2 is located is relatively stable, and the reaction rate oftoluene alkylation can be greatly increased.

As an embodiment of the present application, the method for producingpara-xylene and co-producing light olefins from methanol and toluenecomprises at least the following steps:

(1) passing the mixture of methanol and toluene from the first gasdistributor into the reaction zone of the fluidized bed reactor;

(2) passing methanol from the second gas distributor into the reactionzone of the fluidized bed reactor, the mass ratio of the methanolentering from the second gas distributor to the methanol entering fromthe first gas distributor is in a range from 1:1 to 10:1;

(3) the reactants in the reaction zone are contacted with the catalystto form a gas phase stream comprising para-xylene and light olefins;

(4) passing the gas phase stream into a settling zone, a gas-solidseparator, and entering a subsequent separation section via a productoutlet, and after separation, to obtain ethylene, propylene, butene,para-xylene, dimethyl ether, chain hydrocarbon by-products, aromaticby-products and converted methanol and toluene, the chain hydrocarbonby-products include methane, ethane, propane, butane and C₅₊ chainhydrocarbons, and the aromatic by-products include benzene,ethylbenzene, o-xylene, m-xylene and C₉₊ aromatic hydrocarbons;

(5) returning dimethyl ether and unconverted methanol as raw material tothe fluidized bed reactor via the second gas distributor to recycle, andreturning the aromatic by-products and unconverted toluene as rawmaterial to the fluidized bed reactor via the first gas distributor torecycle;

(6) forming a spent catalyst from the catalyst after carbon depositionin the reaction zone, subjecting the spent catalyst to stripping andregenerating in a fluidized bed regenerator, and passing the regeneratedcatalyst into the fluidized bed reactor via the regenerated catalystdelivery pipe.

In the above method, the fluidized bed reactor has a gas phase linearvelocity ranging from 0.2 m/s to 2 m/s and a temperature ranging from300° C. to 600° C., and the fluidized bed regenerator has a gas phaselinear velocity ranging from 0.2 m/s to 2 m/s and a temperature rangingfrom 500° C. to 800° C.

Example 1

In the fluidized bed reactor as shown in FIG. 1, para-xylene and lightolefins are produced. The fluidized bed reactor comprises the first gasdistributor 1, the second gas distributor 2, the reaction zone 3, thesettling zone 4, the gas-solid separator 5, the stripping zone 6 and theregenerated catalyst delivery pipe 7. The first gas distributor 1 isplaced at the bottom of the reaction zone 3. The second gas distributor2 is placed in the reaction zone 3. The settling zone 4 is above thereaction zone 3. The gas-solid separator 5 is disposed within thesettling zone 4, and the product outlet is set on the top. The strippingzone 6 is below the reaction zone 3, and the upper portion of thereaction zone 3 is connected with the regenerated catalyst delivery pipe7.

The first gas distributor 1 is a branched pipe distributor, and thesecond gas distributor 2 is a microporous gas distributor.

As shown in FIG. 3, the microporous gas distributor includes the intakepipe 2-1, the intake ring pipe 2-2 and the microporous pipe 2-3. Asshown in FIG. 4, the intake pipe 2-1 is connected to two intake ringpipes 2-2, the intake ring pipes 2-2 are uniformly provided with twelvemicroporous pipes 2-3. The microporous pipes 2-3 is a powder metallurgymicroporous pipe, and the spacing between microporous pipes 2-3 is in arange from 150 mm to 200 mm, the pore diameter of the micropores is in arange from 1 μm to 10 μm, the porosity is 35%, and the gas velocity inthe pipe is 5 m/s.

The catalyst in the fluidized bed reactor is a ZSM-5 molecular sievecatalyst.

Material Stream A: the mixture of toluene, aromatic by-products andmethanol. Material stream A enters the reaction zone 3 of the fluidizedbed reactor via the first gas distributor 1, and the mass content ofmethanol in the mixture of material stream A is 2%.

Material Stream B: methanol. The material stream B enters the reactionzone 3 of the fluidized bed reactor from the second gas distributor 2,and the mass ratio of the methanol entering from the second gasdistributor 2 to the methanol entering from the first gas distributor 1is 9:1. The fluidized bed reactor has a gas phase linear velocityranging from 0.8 to 1.0 m/s and a temperature of 450° C. The reactantsin the reaction zone 3 are contacted with the catalyst to form a gasphase stream comprising para-xylene and light olefins. The gas phasestream enters the settling zone 4, the gas-solid separator 5, and entersa subsequent separation section via the product outlet. The catalystforms the spent catalyst after carbon deposition in the reaction zone,and the spent catalyst is fed into the fluidized bed regenerator forregeneration. The gas phase linear velocity of the fluidized bedregenerator is 1.0 m/s and the temperature is 650° C. The regeneratedcatalyst enters the fluidized bed reactor via the regenerated catalystdelivery pipe 7.

The product composition is analyzed by gas chromatography. The resultsshow that the conversion rate of toluene is 41%, the conversion rate ofmethanol is 99.6%, and the mass single-pass yield of para-xylene basedon aromatics is 38%, the selectivity of para-xylene in the xylene isomerin the products is 99%, and the selectivity of(ethylene+propylene+butene) in C₁˜C₆ chain hydrocarbon component is 75%.

Example 2

In the fluidized bed reactor as shown in FIG. 1, para-xylene and lightolefins are produced. The fluidized bed reactor comprises the first gasdistributor 1, the second gas distributor 2, the reaction zone 3, thesettling zone 4, the gas-solid separator 5, the stripping zone 6 and theregenerated catalyst delivery pipe 7. The first gas distributor 1 isplaced at the bottom of the reaction zone 3. The second gas distributor2 is placed in the reaction zone 3. The settling zone 4 is above thereaction zone 3. The gas-solid separator 5 is disposed within thesettling zone 4, and the product outlet is set on the top. The strippingzone 6 is below the reaction zone 3, and the upper portion of thereaction zone 3 is connected with the regenerated catalyst delivery pipe7.

The first gas distributor 1 is a branched pipe distributor, and thesecond gas distributor 2 is a microporous gas distributor.

As shown in FIG. 3, the microporous gas distributor includes the intakepipe 2-1, the intake ring pipe 2-2 and the microporous pipe 2-3. Asshown in FIG. 4, the intake pipe 2-1 is connected to two intake ringpipes 2-2, the intake ring pipes 2-2 are uniformly provided with twelvemicroporous pipes 2-3. The microporous pipes 2-3 is a ceramicmicroporous pipe, and the spacing between microporous pipes 2-3 is in arange from 150 mm to 200 mm, the pore diameter of the micropores is in arange from 20 μm to 40 μm, the porosity is 45%, and the gas velocity inthe pipe is 4 m/s.

The catalyst in the fluidized bed reactor is a ZSM-5 molecular sievecatalyst.

Material Stream A: the mixture of toluene, aromatic by-products andmethanol. Material stream A enters the reaction zone 3 of the fluidizedbed reactor via the first gas distributor 1, and the mass content ofmethanol in the mixture of material stream A is 5%.

Material Stream B: methanol. The material stream B enters the reactionzone 3 of the fluidized bed reactor from the second gas distributor 2,and the mass ratio of the methanol entering from the second gasdistributor 2 to the methanol entering from the first gas distributor 1is 8:1. The fluidized bed reactor has a gas phase linear velocityranging from 1.3 m/s to 1.5 m/s and a temperature of 500° C. Thereactants in the reaction zone 3 are contacted with the catalyst to forma gas phase stream comprising para-xylene and light olefins. The gasphase stream enters the settling zone 4, the gas-solid separator 5, andenters a subsequent separation section via the product outlet. Thecatalyst forms the spent catalyst after carbon deposition in thereaction zone, and the spent catalyst is fed into the fluidized bedregenerator for regeneration. The gas phase linear velocity of thefluidized bed regenerator is 1.0 m/s and the temperature is 600° C. Theregenerated catalyst enters the fluidized bed reactor via theregenerated catalyst delivery pipe 7.

The product composition is analyzed by gas chromatography. The resultsshow that the conversion rate of toluene is 45%, the conversion rate ofmethanol is 98%, and the mass single-pass yield of para-xylene based onaromatics is 45%, the selectivity of para-xylene in the xylene isomer inthe products is 98%, and the selectivity of (ethylene+propylene+butene)in C₁˜C₆ chain hydrocarbon component is 70%.

Example 3

In the fluidized bed reactor as shown in FIG. 1, para-xylene and lightolefins are produced. The fluidized bed reactor comprises the first gasdistributor 1, the second gas distributor 2, the reaction zone 3, thesettling zone 4, the gas-solid separator 5, the stripping zone 6 and theregenerated catalyst delivery pipe 7. The first gas distributor 1 isplaced at the bottom of the reaction zone 3. The second gas distributor2 is placed in the reaction zone 3. The settling zone 4 is above thereaction zone 3. The gas-solid separator 5 is disposed within thesettling zone 4, and the product outlet is set on the top. The strippingzone 6 is below the reaction zone 3, and the upper portion of thereaction zone 3 is connected with the regenerated catalyst delivery pipe7.

The first gas distributor 1 is a plate distributor with blast caps, andthe second gas distributor 2 is a microporous gas distributor.

As shown in FIG. 3, the microporous gas distributor includes the intakepipe 2-1, the intake ring pipe 2-2 and the microporous pipe 2-3. Asshown in FIG. 4, the intake pipe 2-1 is connected to two intake ringpipes 2-2, the intake ring pipes 2-2 are uniformly provided with twelvemicroporous pipes 2-3. The microporous pipes 2-3 is a ceramicmicroporous pipe, and the spacing between microporous pipes 2-3 is in arange from 150 mm to 200 mm, the pore diameter of the micropores is in arange from 5 μm to 20 μm, the porosity is 45%, and the gas velocity inthe pipe is 6 m/s.

The catalyst in the fluidized bed reactor is a ZSM-5 molecular sievecatalyst.

Material Stream A: the mixture of toluene, aromatic by-products andmethanol. Material stream A enters the reaction zone 3 of the fluidizedbed reactor via the first gas distributor 1, and the mass content ofmethanol in the mixture of material stream A is 10%.

Material Stream B: methanol. The material stream B enters the reactionzone 3 of the fluidized bed reactor from the second gas distributor 2,and the mass ratio of the methanol entering from the second gasdistributor 2 to the methanol entering from the first gas distributor 1is 5:1. The fluidized bed reactor has a gas phase linear velocityranging from 0.2 to 0.3 m/s and a temperature of 550° C. The reactantsin the reaction zone 3 are contacted with the catalyst to form a gasphase stream C comprising para-xylene and light olefins. The gas phasestream enters the settling zone 4, the gas-solid separator 5, and entersa subsequent separation section via the product outlet. The catalystforms the spent catalyst after carbon deposition in the reaction zone,and the spent catalyst is fed into the fluidized bed regenerator forregeneration. The gas phase linear velocity of the fluidized bedregenerator is 1.0 m/s and the temperature is 700° C. The regeneratedcatalyst enters the fluidized bed reactor via the regenerated catalystdelivery pipe 7.

The product composition is analyzed by gas chromatography. The resultsshow that the conversion rate of toluene is 46%, the conversion rate ofmethanol is 96%, and the mass single-pass yield of para-xylene based onaromatics is 48%, the selectivity of para-xylene in the xylene isomer inthe products is 97%, and the selectivity of (ethylene+propylene+butene)in C₁˜C₆ chain hydrocarbon component is 74%.

Example 4

In the fluidized bed reactor as shown in FIG. 2, para-xylene and lightolefins are produced. The fluidized bed reactor comprises the first gasdistributor 1, the second gas distributor 2, the reaction zone 3, thesettling zone 4, the gas-solid separator 5, the stripping zone 6 and theregenerated catalyst delivery pipe 7. The first gas distributor 1 isplaced at the bottom of the reaction zone 3. The second gas distributor2 is placed in the reaction zone 3. The settling zone 4 is above thereaction zone 3. The gas-solid separator 5 is disposed within thesettling zone 4, and the product outlet is set on the top. The strippingzone 6 is below the reaction zone 3, and the upper portion of thereaction zone 3 is connected with the regenerated catalyst delivery pipe7.

The first gas distributor 1 is a branched pipe distributor, and thesecond gas distributor 2 is a microporous gas distributor.

As shown in FIG. 3, the microporous gas distributor includes the intakepipe 2-1, the intake ring pipe 2-2 and the microporous pipe 2-3. Asshown in FIG. 4, the intake pipe 2-1 is connected to two intake ringpipes 2-2, the intake ring pipes 2-2 are uniformly provided with twelvemicroporous pipes 2-3. The microporous pipes 2-3 is a ceramicmicroporous pipe, and the spacing between microporous pipes is in arange from 150 mm to 200 mm, the pore diameter of the micropores is in arange from 5 μm to 20 μm, the porosity is 45%, and the gas velocity inthe pipe is 6 m/s.

The catalyst in the fluidized bed reactor is a ZSM-5 molecular sievecatalyst.

Material Stream A: the mixture of toluene, aromatic by-products andmethanol. Material stream A enters the reaction zone 3 of the fluidizedbed reactor via the first gas distributor 1, and the mass content ofmethanol in the mixture of material stream A is 20%.

Material Stream B: methanol. The material stream B enters the reactionzone 3 of the fluidized bed reactor from the second gas distributor 2,and the mass ratio of the methanol entering from the second gasdistributor 2 to the methanol entering from the first gas distributor 1is 4:1. The fluidized bed reactor has a gas phase linear velocityranging from 1.5 m/s to 2.0 m/s and a temperature of 530° C. Thereactants in the reaction zone 3 are contacted with the catalyst to forma gas phase stream comprising para-xylene and light olefins. The gasphase stream enters the settling zone 4, the gas-solid separator 5, andenters a subsequent separation section via the product outlet. Thecatalyst forms the spent catalyst after carbon deposition in thereaction zone, and the spent catalyst is fed into the fluidized bedregenerator for regeneration. The gas phase linear velocity of thefluidized bed regenerator is 2.0 m/s and the temperature is 700° C. Theregenerated catalyst enters the fluidized bed reactor via theregenerated catalyst delivery pipe 7.

The product composition is analyzed by gas chromatography. The resultsshow that the conversion rate of toluene is 49%, the conversion rate ofmethanol is 91%, and the mass single-pass yield of para-xylene based onaromatics is 51%, the selectivity of para-xylene in the xylene isomer inthe products is 95%, and the selectivity of (ethylene+propylene+butene)in C₁˜C₆ chain hydrocarbon component is 71%.

While the present application has been described above with reference topreferred embodiments, but these embodiments are not intended to limitthe claims. Without departing from the spirit of the presentapplication, people skilled in the art will be able to make severalpossible variations and modifications and thus the protection scopeshall be determined by the scope as defined in the claims.

1. A fluidized bed gas distributor, comprising a first distributor and asecond distributor, wherein the first distributor is located at thebottom of the fluidized bed, and the second distributor is located in atleast one region of the gas flow downstream from the first distributor.2. The fluidized bed gas distributor of claim 1, wherein the seconddistributor comprises an intake pipe, a microporous pipe and an intakering pipe; the intake pipe is connected with the microporous pipe, thegas is introduced by the intake pipe from the outside of the fluidizedbed into the microporous pipe in the fluidized bed; the intake ring pipeis connected with the intake pipe, the intake ring pipe is disposed on aplane perpendicular to the flow direction of the gas from the firstdistributor; the microporous pipe is disposed on the intake ring pipeand perpendicular to a plane of the intake ring pipe.
 3. The fluidizedbed gas distributor of claim 1, wherein material stream A enters thefluidized bed through the first distributor, material stream B entersthe fluidized bed through the second distributor and contacts with atleast a portion of the gas of the material stream A.
 4. The fluidizedbed gas distributor of claim 1, wherein the first distributor is atwo-dimensional gas distributor and the first distributor distributesthe gas on the plane in which the first distributor is located at thebottom of the fluidized bed.
 5. The fluidized bed gas distributor ofclaim 1, wherein the second distributor is a three-dimensional gasdistributor and the second distributor distributes the gas in at least aportion of the reaction space in which the second distributor is locatedin the fluidized bed.
 6. The fluidized bed gas distributor of claim 1,wherein the first distributor is a branched pipe distributor and/or aplate distributor with blast caps.
 7. The fluidized bed gas distributorof claim 2, wherein the microporous pipe is a ceramic microporous pipeand/or a metal microporous sintered pipe.
 8. The fluidized bed gasdistributor of claim 2, wherein the side and end faces of themicroporous pipe have micropores with a pore diameter ranging from 0.5μm to 50 μm and a porosity ranging from 25% to 50%, the gas velocity inthe pipe of the microporous pipe is in a range from 0.1 m/s to 10 m/s;preferably, the gas velocity in the pipe of the microporous pipe is in arange from 1 m/s to 10 m/s.
 9. (canceled)
 10. The fluidized bed gasdistributor of claim 2, wherein the microporous pipes are arranged inparallel with each other; the intake ring pipes are arranged in aconcentric ring or planar spiral in the same plane.
 11. The fluidizedbed gas distributor of claim 1, wherein the fluidized bed gasdistributor is used in a fluidized bed reactor for producing para-xyleneand co-producing light olefins.
 12. A fluidized bed reactor, comprisingat least one of the fluidized bed gas distributors according to claim 1.13. The fluidized bed reactor of claim 12, wherein the fluidized bedreactor comprises a reaction zone, a settling zone, a gas-solidseparator, a stripping zone and a regenerated catalyst delivery pipe;the first distributor is placed at the bottom of the reaction zone, thesecond distributor is placed above the first distributor, the settlingzone is above the reaction zone, the settling zone is provided with thegas-solid separator, the stripping zone is below the reaction zone, andthe regenerated catalyst delivery pipe is connected with the reactionzone; preferably, the regenerated catalyst delivery pipe is connectedwith the upper portion of the reaction zone; preferably, the regeneratedcatalyst delivery pipe is connected with the bottom of the reactionzone.
 14. (canceled)
 15. (canceled)
 16. A method for producingpara-xylene and co-producing light olefins from methanol and/or dimethylether and toluene, wherein at least one of the fluidized bed gasdistributors according to claim 1 is used; the method for producingpara-xylene and co-producing light olefins comprises at least thefollowing steps: (1) passing the material stream A from the firstdistributor into the reaction zone of the fluidized bed reactor, thereaction zone containing a catalyst; the material stream A containingtoluene, or the material stream A containing methanol and/or dimethylether and toluene; (2) passing the material stream B containing methanoland/or dimethyl ether from the second distributor into the reaction zoneof the fluidized bed reactor; (3) in the reaction zone, contactingmethanol and/or dimethyl ether and toluene from the material stream Aand/or the material stream B, with the catalyst to form material streamC comprising para-xylene and light olefins.
 17. The method of claim 16,wherein the method for producing para-xylene and co-producing lightolefins from methanol and/or dimethyl ether and toluene furthercomprises the following steps: (4) passing the material stream C intothe settling zone and the gas-solid separator to separate the materialstream C to obtain light olefins, para-xylene, chain hydrocarbonby-products, aromatic by-products and unconverted toluene, unconvertedmethanol and/or dimethyl ether; (5) returning unconverted methanoland/or dimethyl ether to the fluidized bed reactor via the seconddistributor; returning the aromatic by-products and unconverted tolueneto the fluidized bed reactor via the first distributor; (6) forming aspent catalyst from the catalyst after carbon deposition in the reactionzone, the spent catalyst is then stripped in a stripper and regenerated,in a regenerator to obtain a regenerated catalyst; passing theregenerated catalyst into the fluidized bed reactor via the regeneratedcatalyst delivery pipe; preferably, the mass ratio of methanol and/ordimethyl ether in material stream B to methanol and/or dimethyl ether inmaterial stream A is in a range from 1:1 to 10:1.
 18. (canceled)
 19. Themethod of claim 16, wherein the sum of the mass percentages of methanoland dimethyl ether in material stream A is in a range from 0% to 30%;preferably, the sum of the mass percentages of methanol and dimethylether in material stream A is in a range from 2% to 20%.
 20. (canceled)21. The method of claim 16, wherein the fluidized bed reactor has a gasphase linear velocity ranging from 0.2 m/s to 2 m/s and a reactiontemperature ranging from 300° C. to 600° C.; preferably, the regeneratorhas a gas phase linear velocity ranging from 0.2 m/s to 2 m/s and aregeneration temperature ranging from 500° C. to 800° C.
 22. (canceled)23. A method for producing para-xylene and co-producing light olefinsfrom methanol and/or dimethyl ether and toluene, wherein at least one ofthe fluidized bed reactors according to claim 12 is used; the method forproducing para-xylene and co-producing light olefins comprises at leastthe following steps: (1) passing the material stream A from the firstdistributor into the reaction zone of the fluidized bed reactor, thereaction zone containing a catalyst; the material stream A containingtoluene, or the material stream A containing methanol and/or dimethylether and toluene; (2) passing the material stream B containing methanoland/or dimethyl ether from the second distributor into the reaction zoneof the fluidized bed reactor; (3) in the reaction zone, contactingmethanol and/or dimethyl ether and toluene from the material stream Aand/or the material stream B, with the catalyst to form material streamC comprising para-xylene and light olefins.
 24. The method of claim 23,wherein the method for producing para-xylene and co-producing lightolefins from methanol and/or dimethyl ether and toluene furthercomprises the following steps: (4) passing the material stream C intothe settling zone and the gas-solid separator to separate the materialstream C to obtain light olefins, para-xylene, chain hydrocarbonby-products, aromatic by-products and unconverted toluene, unconvertedmethanol and/or dimethyl ether; (5) returning unconverted methanoland/or dimethyl ether to the fluidized bed reactor via the seconddistributor; returning the aromatic by-products and unconverted tolueneto the fluidized bed reactor via the first distributor; (6) forming aspent catalyst from the catalyst after carbon deposition in the reactionzone, the spent catalyst is then stripped in a stripper and regenerated,in a regenerator to obtain a regenerated catalyst; passing theregenerated catalyst into the fluidized bed reactor via the regeneratedcatalyst delivery pipe; preferably, the mass ratio of methanol and/ordimethyl ether in material stream B to methanol and/or dimethyl ether inmaterial stream A is in a range from 1:1 to 10:1.
 25. The method ofclaim 24, wherein the sum of the mass percentages of methanol anddimethyl ether in material stream A is in a range from 0% to 30%;preferably, the sum of the mass percentages of methanol and dimethylether in material stream A is in a range from 2% to 20%.
 26. The methodof claim 24, wherein the fluidized bed reactor has a gas phase linearvelocity ranging from 0.2 m/s to 2 m/s and a reaction temperatureranging from 300° C. to 600° C.